Synthetic fuels and chemicals production with in-situ CO2 capture

ABSTRACT

Novel redox based systems for fuel and chemical production with in-situ CO 2  capture are provided. A redox system using one or more chemical intermediates is utilized in conjunction with liquid fuel generation via indirect Fischer-Tropsch synthesis, direct hydro genation, or pyrolysis. The redox system is used to generate a hydrogen rich stream and/or CO 2  and/or heat for liquid fuel and chemical production. A portion of the byproduct fuels and/or steam from liquid fuel and chemical synthesis is used as part of the feedstock for the redox system.

CROSS-REFERENCE TO RELATED APPLICATION(S)

This is a continuation of U.S. patent application Ser. No. 13/394,396,filed on Mar. 6, 2012, which is a U.S. national stage entry ofInternational Patent Application No. PCT/US2010/048121, filed on Sep. 8,2010, which claims priority to U.S. Provisional Patent Application No.61/240,446, filed on Sep. 8, 2009, the entire contents of all of whichare fully incorporated herein by reference.

The present invention is generally directed to systems and methods forsynthetic fuels and chemical products generation with in-situ CO₂capture. A reduction-oxidation (redox) system using one or more chemicalintermediates is generally utilized in conjunction with liquid fuelgeneration via indirect CO₂ hydrogenation, direct hydrogenation, orpyrolysis.

Fossil fuels including crude oil, natural gas, and coal provide morethan 85% of today's energy supply. These fossil fuels are usuallytransformed to carriers such as electricity and liquid transportationfuels prior to utilization by end consumers. Electricity is mainlyproduced by relatively abundant energy sources such as coal, naturalgas, and nuclear. In contrast, liquid transportation fuel is almostexclusively obtained from crude oil, whose supply is relatively insecurewith volatile prices. With an increasing energy demand and concomitantconcerns over carbon emissions from fossil fuel usage, affordablesynthetic transportation fuels from more abundant resources such ascoal, biomass, and oil shale are desirable. To address the environmentalconcerns, the next generation synthetic fuel production processes needto be able to capture pollutants generated in the process. Thesepollutants include CO₂, sulfur compounds, and mercury, among others.

Synthetic fuel is generated from gaseous fuels such as natural gasthrough reforming and the Fischer-Tropsch (“F-T”) scheme. Solid fuelssuch as coal, biomass, and pet coke can be converted to synthetic fuelthrough indirect liquefaction (gasification—water gasshift—Fischer-Tropsch), direct liquefaction, or pyrolysis. These systemsare, however, more capital intensive than oil refining processes.Moreover, their energy conversion efficiencies are relatively low.

Synthetic fuel can also be generated from biomass via biochemicalroutes. However, a large amount of process water is utilized. Moreover,the biochemical approaches have stringent requirements on the feedstock.

All the aforementioned processes involve CO₂ emissions. CO₂ capture fromthese processes associates with notable energy losses and hencedecreases in process efficiency.

Embodiments of the present invention provide alternatives to producesynthetic fuel from naturally occurring carbonaceous fuel sources withhigh efficiency and effective CO₂ capture.

Embodiments of the present invention are generally directed to novelredox based systems for fuel and chemical production with in-situ CO₂capture. A redox system using one or more chemical intermediates isgenerally utilized in conjunction with liquid fuel generation viaindirect Fischer-Tropsch synthesis, direct hydrogenation, or pyrolysis.The redox system is used to generate a hydrogen rich stream and/or CO₂and/or heat for liquid fuel and chemical production. A portion of thebyproduct fuels and/or steam from liquid fuel and chemical synthesis isused as part of the feedstock for the redox system.

Additional features and advantages provided by embodiments of thepresent invention will be more fully understood in view of the followingdetailed description.

The following detailed description of the illustrative embodiments ofthe present invention can be best understood when read in conjunctionwith the following drawings, where like structure is indicated with likereference numerals and in which:

FIG. 1 illustrates a synthetic liquid fuel production embodiment thatutilizes a combination of indirect reforming/gasification ofcarbonaceous feedstock and Fischer-Tropsch synthesis.

FIG. 2 is a schematic illustration of another embodiment illustratingthe integration of the indirect reforming/gasification andFischer-Tropsch synthesis.

FIG. 3 illustrates another embodiment of the integration of an ironoxide based gaseous fuel indirect reforming/gasification system andFischer-Tropsch synthesis. Coal and a coal gasification unit are used inthis case to produce syngas fuel. Methane and hydrocarbons can also bedirectly used in this system. Alternatively, a reformer can be installedin place of the gasification unit (gasifier) to convert hydrocarbonfuels.

FIG. 4 illustrates another embodiment using the integration of an ironoxide based solid fuel indirect gasification system and Fischer-Tropschsynthesis. Besides biomass and coal, other solid fuels such as pet coke,tar sands, oil shale, and waste derived fuel can also be used in thissystem.

FIG. 5 illustrates another embodiment using the integration of a sorbentenhanced reforming/water gas shift system and Fischer-Tropsch synthesis.Gaseous fuels such as syngas and light hydrocarbons can be used in thissystem.

FIG. 6 is a schematic of another embodiment showing the integrationbetween a direct coal to liquid sub-system and an indirect carbonaceousfuel reforming/gasification sub-system. A sorbent enhancedreforming/water gas shift system can also be used to replace the redoxbased indirect reforming/gasification sub-system.

FIG. 7 shows another embodiment of the integration between a biomasspyrolyzer and an indirect carbonaceous fuel reforming/gasificationsub-system for bio-oil synthesis.

FIG. 8 is another embodiment illustrating the integration scheme betweena biomass pyrolyzer and an indirect carbonaceous fuelreforming/gasification sub-system for bio-oil synthesis.

FIG. 9 illustrates additional reducer designs for pulverizedcoal/biomass conversion in a countercurrent moving bed with coal/biomasspowder flowing upwards and metal oxide composites flowing downwards.

FIG. 9. (a) illustrates a moving bed reducer design for pulverized coaland biomass conversion; FIG. 9 (b) illustrates a potential design forcoal injection and conversion.

Embodiment of the present invention are generally directed to systemsand methods for converting carbonaceous fuels into synthetic fuels withminimal carbon emission and improved energy conversion efficiency. Suchsystems and methods generally include an indirect fuelreforming/gasification sub-system and a liquid fuel synthesissub-system.

Based on the technique through which the synthetic fuel is produced, thevarious embodiments of the present invention can be generally groupedinto three categories, i.e. indirect synthetic fuel generationintegrated with an indirect fuel reforming/gasification sub-system,direct synthetic fuel generation integrated with an indirectreforming/gasification sub-system, and direct pyrolysis systemintegrated with an indirect fuel combustion sub-system. The followingspecification discusses the three categories respectively.

The indirect synthetic fuel generation system, which is strategicallyintegrated with an indirect fuel reforming/gasification sub-system, isgenerally represented by FIGS. 1-5.

The indirect conversion of carbonaceous fuels such as coal and naturalgas to synthetic liquid fuel through gasification/reforming followed byFischer-Tropsch synthesis is well established. The processes, however,are inefficient due to the large irreversibility of thegasification/reforming step and the highly exothermic nature of theFischer-Tropsch synthesis reactions and the inefficiency associated withthe heat recovery and utilization. Further, significant energy losseswill be incurred if the carbon generated in the process is captured. Inaddition, the indirect synthetic fuel generation systems are highlycapital intensive.

The increasing concerns over energy security and CO₂ emissions have castserious doubt on both the environmental and economical acceptability ofindirect synthetic fuel generation systems. To reduce the cost andcarbon footprint of the indirect liquid fuel synthesis systems, drasticimprovement in process energy conversion efficiencies coupled with CO₂capture are highly desirable. Embodiments of the present inventionstrategically integrate an indirect gasification/reforming sub-systemwith Fischer-Tropsch sub-system to achieve effects that: 1) reduce theirreversibility of the overall synthetic fuel product system; 2) improvethe energy conversion efficiency; and 3) capture the CO₂ generated inthe process.

According to one aspect, carbonaceous fuel such as coal, biomass, petcoke, syngas, natural gas, extra heavy oil, wax, and oil shale, arefirst converted into separate streams of CO₂ and H₂ through theassistance of one or more chemical intermediates. The H₂ and a portionof the CO₂ are then reacted in a Fischer-Tropsch synthesis reactor toproduce synthetic fuels and chemicals. The remaining CO₂ is obtained ina concentrated form and can be readily sequestrated. The conversion ofCO₂ and H₂, as opposed to CO and H₂, in the Fischer-Tropsch reactorreduces the exothermicity of the F-T reaction. Moreover, this schemepotentially reduces the endothermicity of the gasification/reformingstep. As a result, the overall process irreversibility can be reduced.Moreover, the steam produced from the exothermic F-T reactor is readilyavailable for hydrogen generation in the gasification/reformingsub-system. While the use of CO₂ and H₂ for F-T synthesis was studied inthe 1990s, the method for CO₂ and H₂ generation from carbonaceous fuelsand the unique integration schemes between the CO₂/II₂ generationsub-system described herein are novel.

FIG. 1 is generally directed to an integration scheme of a redox basedgasification/reforming sub-system and an F-T sub-system. With thisconfiguration, a carbonaceous fuel is indirectly gasified/reformed intotwo separate streams of CO₂ and H₂. The two streams are then cooled andintroduced into the F-T sub-system to produce liquid fuels. Thereactions, which are not balanced, in this process include:MeO_(x)+C_(x)H_(y)O_(z)→CO₂+H₂O+MeO_(y)(Reactor 1)MeO_(y)+H₂O→MeO_(z)+H₂(Reactor 2, y<z≤x)MeO_(z)+O₂→MeO_(z)(Reactor 3, optional)CO₂+H₂→—(CH₂)—+H₂O(CO₂ hydrogenation)Here C_(x)H_(y)O_(z), refers to a carbonaceous fuel in general. Me is ametal or metal mixture that can be reduced by the carbonaceous fuel andsubsequently oxidized by steam and air. Such metals include Fe, Co, In,Mn, Sn, Zn, Cu, W, and combinations thereof.

Reactor 1 is typically operated at 400-1200° C. and 1.01×10⁵ Pa−8.10×10⁶Pa (1-80 atm). Reactor 2 is operated at a temperature of 0-300° C. lowerthan Reactor 1. Reactor 3, which is optional depending on the type ofmetal and the system configuration, is operated at a temperature 0-400°C. higher than Reactor 1. In preferred embodiments, Reactor 1 isoperated at 600-900° C. The gasification/reforming sub-system isoperated at 1.01×10⁵ Pa−3.04×10⁶ Pa (1-30 atm).

In certain embodiments, Reactor 1 is endothermic. A portion of thereduced solids from Reactor 1 is directly sent to Reactor 3 foroxidation with oxygen containing gas. The heat released in Reactor 3 isused to compensate for the heat required in Reactor 1. The extra heatgenerated in Reactor 3 is used for power generation to support theparasitic power usage. A small portion of the hydrogen from Reactor 2can be used for fuel product upgrading.

As showing in FIG. 1, carbonaceous fuel is fed near the bottom ofReactor 1. In one embodiment, the carbonaceous fuel comprises solidparticles which are suspended buy the gases in a lower tapered sectionof Reactor 1 until they are at least to 50% converted before beingelutriated towards the tope of Reactor 1. CO₂ rich gas and H₂ rich gasare produced from Reactor 1 and Reactor 2, respectively. These gaseousstreams, which may contain steam, can be condensed prior to F-Tsynthesis. Alternatively, these gaseous streams can be directly used forF-T synthesis.

The F-T sub-system is operated at 200-500° C. and 1.01×10⁶ Pa−8.10×10⁷Pa (10-100 atm). In some embodiments, compression of the CO₂ rich gasand H₂ rich gas from the gasification/reforming sub-system arecompressed.

Sulfur may present in the carbonaceous fuel, contaminating the CO₂ richgas and H₂ rich gas streams. One or more sulfur removal units may beused to clean up the product gas streams. In the case where an ironbased catalyst is used for F-T synthesis, a high temperature sorbent bedusing solid sorbents such as CaO, ZnO, etc. can be used to reduce thesulfur contaminants to levels of 100 ppm or less. When a less sulfurtolerant catalyst such as cobalt based F-T catalyst is used for F-Tsynthesis, additional sulfur removal steps such as that using MDEA,SELEXOL (trade name), or Rectisol (trade name) may be used. In the casewhen low sulfur fuel such as low sulfur biomass and sulfur free naturalgas or syngas is used, the sulfur removal units are not necessary.

FIG. 2 illustrates another process configuration which integrates theredox based gasification/reforming sub-system and the F-T sub-system. Inthis configuration, the unconverted fuels from the F-T sub-system arerecycled back to Reactor 1 along with the carbonaceous fuel feedstock.By doing so, the byproduct from the F-T sub-system is converted to H₂and CO₂, increasing the liquid fuel yield and selectivity of theprocess. In addition, the steam generated from the F-T sub-system isredirected to Reactor 2 of the gasification/reforming sub-system,reducing the need for steam generation in the process. The strategicutilization of the products and byproducts of both F-T andgasification/reforming sub-systems and their integration-recirculationschemes reduce the exergy loss of the overall process while increasingthe yield of desired product, either chemical or synthetic liquid fuel.Any CO₂ generated in the process is readily sequestrable. As a result,the process is significantly less carbon intensive and more efficientthan conventional coal to liquids schemes.

FIG. 3 further illustrates a more detailed process configuration,integrating an iron oxide based gasification/reforming sub-system and anF-T sub-system. In this embodiment, the gasification/reformingsub-system comprises a gasification/reforming unit and an iron basedredox unit. Solid fuel is first converted into a gaseous fuel mixture.The gaseous fuel is then injected to the reducer of the iron oxide redoxsystem for hydrogen and CO₂ generation. A hot gas cleanup system may berequired where the gaseous fuel is contaminated with a high level ofsulfur. The three reactor iron oxide based redox system is used toconvert the fuel in a manner similar to that disclosed in Thomas U.S.Pat. No. 7,767,191; Fan PCT Application No. WO 2007082089; and Fan PCTApplication No. WO 2010037011. The first reactor, the reducer, isconfigured to oxidize the carbonaceous fuel into CO₂ and steam whilereducing a metal oxide based oxygen carrier, such that the averagevalence of the metal is less than 1. The heat required or generated inthe reducer is provided or removed by the oxygen carrier particle. Thesecond reactor, the oxidizer, is configured to (partially) oxidize aportion of the reduced oxygen carrier with steam. The third reactor, thecombustor, combusts the partially oxidized oxygen carrier from theoxidizer and the remaining portion of the reduced oxygen carrier fromthe reducer with air. The reactions in the iron oxide redox systeminclude, without balancing the equations:Fe₂O₃+Fuel→Fe/FeO+CO₂+H₂O (avg. valence of Fe is <1) (Reducer)Fe/FeO+H₂O→Fe₃O₄+H₂ (Oxidizer)Fe₃O₄+O₂ (Air)→Fe₂O₃ (Combustor)

In one embodiment, all of the hydrogen from the oxidizer and a portionof the CO₂ from the reducer are introduced to the Fischer-Tropschreactor to generate a mixture of hydrocarbons. The hydrocarbon mixtureis then separated and refined. The fraction of the fuel mixture of lowereconomic value, e.g. unconverted syngas, light hydrocarbons, andnaphtha, is sent to either the reducer or the gasifier/reformer toenhance carbon utilization. In essence, most of the carbon in the fuelis either fixed in the final synthetic fuel product or in theconcentrated CO₂ stream which is ready for sequestration after moderatecompression. Hence, the net life cycle CO₂ emissions of the system arecomparable to petroleum based gasoline and diesel when coal is used asthe fuel (with CO₂ capture and sequestration). In the case when biomassand natural gas are used as the fuel, the net life cycle CO₂ emission ismuch lower or even negative. In a carbon constrained scenario, acombination of feedstock such as coal/biomass, coal/natural gas can beused to reduce the CO₂ emissions while taking advantage of abundantlyavailable coal.

The F-T reactor generates a large amount of steam for F-T coolingpurposes, and a portion of the steam is used in the oxidizer forhydrogen generation. The rest of the steam, after supplemental firing orsuperheating with a small portion of byproduct fuel and heat exchangingwith high temperature exhaust gas streams in the process, is used forpower generation to meet the parasitic energy needs.

The oxygen carrier comprises a plurality of ceramic composite particleshaving at least one metal oxide disposed on a support. Ceramic compositeparticles are described in Thomas U.S. Pat. No. 7,767,191; Fan,published PCT Application No. WO 2007082089; and Fan, PCT ApplicationNo. WO 2010037011. In addition to the particles and particle formula andsynthesis methods described in Thomas, applicants, in a furtherembodiment, have developed novel methods and supporting materials toimprove the performance and strength of the ceramic composite particlesused in the present system.

The novel methods include the step of mixing a metal oxide with at leastone ceramic support material in slurry form followed by drying,granulation, and pelletization. Ceramic support materials in addition tothose described in the prior publications include magnesium oxide,bentonite, olivine, kaoline, and sepiolite. Olivine is also used as apromoter for hydrocarbon conversion.

FIG. 4 illustrates an embodiment in which an iron based three reactorredox system directly converts solid fuels into CO₂ and H₂ followed byFischer-Tropsch synthesis. In this embodiment, an iron oxide basedoxygen carrier is reduced by a solid fuel. This is followed by steamregeneration and air combustion in a similar manner as the embodimentshown in FIG. 3.

Referring now to the reduction reaction in the first reactor of FIG. 4,i.e. the reducer, the reducer utilizes various solid carbonaceous fuelssuch as biomass, coal, tars, oil shales, oil sands, tar sand, wax, andcoke to reduce the iron oxide containing ceramic composite to produce amixture of reduced metal and/or metal oxide. In addition to the solidcarbonaceous fuel, the byproducts and unconverted fuel from the liquidfuel synthesis sub-system are also converted in the reducer. Thepossible reduction reactions include:FeO_(x)+Fuel→FeO_(y)+CO₂+H₂OFuel+CO₂→CO+H₂Fuel+H₂O→CO+H₂FeO_(x)+CO/H₂→FeO_(y)+CO₂/H₂OThe preferred overall reaction is:Fe₂O₃+Fuel→Fe/FeO+CO₂+H₂O

Specifically, metallic iron (Fe) is formed in the reducer.Simultaneously, an exhaust stream that contains at least 80% CO₂ (drybasis) is produced from the reducer. In preferred embodiments, the CO₂concentration exceeds 95% and is directly sequestrable.

The preferred designs of the reducer include a moving bed reactor withone or more stages, a multistage fluidized bed reactor, a step reactor,a rotary kiln, or any suitable reactors or vessels known to one ofordinary skill in the art that provide a countercurrent gas-solidcontacting pattern. The counter-current flow pattern between solid andgas is used to enhance the gas and solid conversion. The counter-currentflow pattern minimizes the back-mixing of both solid and gas. Moreover,this flow pattern keeps the solid outlet of the reactor at a morereductive environment while the gas outlet of the reactor in maintainedin a more oxidative environment. As a result, the gas and solidconversion are both enhanced.

Referring back to the oxidation reaction in the second reactor in FIG.4, i.e. the oxidizer, the oxidizer converts a portion of the ironcontaining oxygen carrier particles from the reducer to higher oxidationstate using steam generated from Fischer-Tropsch cooling. The possiblereactions include:Fe+H₂O→FeO+CO/H₂3FeO+H₂O→Fe₃O₄+CO/H₂

The preferred designs of the oxidizer also include a moving bed reactorand other reactor designs that provided a countercurrent gas-solidcontacting pattern. A countercurrent flow pattern is preferred so thathigh steam to hydrogen and CO₂ to CO conversion are achieved.

Referring back to the oxidation reaction in the third reactor in FIG. 4,i.e. the combustor, air or other oxygen containing gas is used tocombust the remaining portion of the reducer solids product and all theoxidizer solids product. The possible reactions in the combustorinclude:Fe/FeO/Fe₃O₄+O₂→Fe₂O₃Alternatively, all the reducer oxygen carrier product will be introducedto the oxidizer to react with a sub-stoichiometric amount of steam.Substantially all of the partially regenerated oxygen carrier from theoxidizer will then be introduced to the combustor. By doing this, noby-pass solids stream is needed.

The preferred reactor designs for the combustor include a fast fluidizedbed reactor, an entrained bed reactor, a transport bed reactor, or amechanical conveying system. The functions of the combustor include:oxidation of the oxygen carrier to a higher oxidation state; andre-circulation of the oxygen carrier to the inlet of the reducer foranother redox cycle.

The combustor is highly exothermic. The heat generated in the combustorcan be used to compensate for the heat required in the reducer. Thisheat can also be used to preheat the feed streams and to generate powerfor parasitic energy consumptions. The high pressure gaseous streamsdischarged from the system can be used to drive expanders for gascompression.

Table 1 illustrates the mass flow of the major streams in a process whenIllinois #6 coal and switchgrass are used as the feedstock and syntheticdiesel is the product. Table 2 illustrates the energy balance of thesystem.

TABLE 1 Mass Balance of the Integrated reforming/gasification - Fischer-Tropsch System for Liquid fuel Synthesis from coal CO₂ from H₂ RichSynthetic Diesel Coal Reducer Stream from from Fuel Production (feed,kg/s) (kmol/s) Oxidizer (kmol/s) Sub-System (bbl/day) 36.9 2.2 4.5 (pureH₂ is 2.9) 8700

TABLE 2 Energy Balance of the Integrated reforming/gasification -Fischer-Tropsch System for Liquid fuel Synthesis from coal Power FuelProcess Coal Parasitic Generation Production Efficiency (MW_(th)) Power(MWe) (MWe) (MW_(th)) (%) 1000 −80 82 620 62.2%

Table 3 illustrates the mass and energy flow of the major streams in aprocess when switchgrass is used as the feedstock and synthetic dieselis the product.

TABLE 3 Mass and Energy Balance of the Integratedreforming/gasification - Fischer-Tropsch System for Liquid fuelSynthesis from switchgrass Synthetic Diesel Process Switchgrass BiomassThermal from Fuel Production Efficiency (Dry feed, kg/s) Input (MW_(th))Sub-System (bbl/day) (%) 5.3 100 818 55.5

Although the cases exemplified by Tables 1-3 are specific to the type offeedstock, product, reforming/gasification sub-system, and liquid fuelproduction system, the choices for the aforementioned parameters have alarge degree of freedom. For instance, multiple types of solids fuelscan be used as the feed and various synthetic fuel products can beproduced.

FIG. 5 illustrates schematically in an embodiment which thereforming/gasification sub-system is comprised of sorbent enhancedreforming/gasification units. In this embodiment, a calcium basedsorbent enhanced reforming process is used as the reforming/watersplitting block. The fuel, which can be carbonaceous feed and/orbyproduct from the liquid fuel synthesis sub-system, is reformed/shiftedto H₂ with the presence of CaO/Ca(OH)₂ sorbent and steam generated fromthe F-T reactor:CaO+Fuel+H₂O→CaCO₃+H₂The spent sorbent is then regenerated at high temperatures using thewaste heat from the system in the calciner:CaCO₃→CaO+CO₂A portion of the byproduct from the liquid fuel synthesis sub-system iscombusted to provide the heat for calcination reaction. A hydration stepis optionally added to reactivate the sorbent. The concentrated CO₂ fromthe calciner is then compressed and sequestered.

The hydrogen and a portion of CO₂ produced from the sorbent enhancedreforming scheme are then used to generate synthetic fuel. Compressionof the CO₂ stream is required prior to fuel synthesis.

FIG. 6 illustrates an embodiment showing the integration between adirect liquefaction sub-system and the reforming/gasificationsub-system. The reforming/gasification sub-system is identical to thoseexemplified in FIG. 1 to FIG. 5, i.e. both metal oxide redox based andsorbent enhanced reforming/gasification sub-systems can be used. Theliquid fuel synthesis sub-system comprises a single or two stage directliquefaction reactor and a refining system. Coal slurry is directlyconverted to hydrocarbons with the presence of catalyst as well ashydrogen from the reforming/gasification sub-system. The pressure of thedirect liquefaction reactor is 5.05×10⁶ Pa−1.01×10⁷ Pa (50-100 atm) andthe temperature is 400-650° C. The light fraction of the fuel and thebyproduct such as heavy residue and char from the refining system areused as the fuel for the reforming/gasification sub-system. Moreover,steam generated in the coal liquefaction unit is also used for hydrogenproduction in the reforming/gasification sub-system. To generalize, theintegrated system uses the byproduct from the liquid fuel synthesissub-system to generate hydrogen for direct coal liquefaction. Moreover,nearly all the carbon, expect for that in the fuel product, is convertedto a CO₂ rich exhaust gas stream from the reforming/gasificationsub-system. The CO₂ rich stream is ready to be sequestered.

FIG. 7 illustrates an embodiment in which there is integration between afast pyrolysis process and a redox based fuel combustion process.Biomass can be converted into bio-oil via a fast pyrolysis process. Fastpyrolysis, however, requires effective control of biomass temperatureand notable heat input. In this embodiment, a metal oxide based two stepredox process is used to provide heat for the pyrolyzer while capturingthe carbon byproduct generated in the process.

The metal oxide is used as the carrier for both oxygen and heat. In thefirst unit, the reducer, high temperature metal oxide (600-1400° C.) isreduced by the residue char and light fractions from the pyrolyzer andrefining block:MeO_(x)+unwanted fuel from pyrolyzer and refining block→MeO_(y)+CO₂This step is mostly endothermic, the hot MeO_(y) exiting the reducer isat a temperature ranging between 400-750° C.

The MeO_(y) from the reducer enters into the prolyzer where it providesheat to the biomass feedstock for fast pyrolysis. The MeO_(y) may becomefurther reduced in the pyrolyzer to MeO_(z). The temperature of theMeO_(z) exiting the pyrolyzer ranges between 300-650° C. The reducer andpyrolyzer can be either a moving bed or a fluidized bed. A fluidized bedis preferred for the pyrolyzer.

The MeO_(z) from the pyrolyzer is then introduced to the oxidizer, whichis similar to the combustor unit described with respect to FIG. 1 toFIG. 4. In the oxidizer, MeO_(z) is combusted with oxygen containing gassuch as air to regenerate to MeO_(x):MeO_(z)+O₂→MeO_(x)The outlet temperature of the oxidizer ranges from 600-1400° C. Thepreferred reactor designs for the oxidizer include a fast fluidized bedreactor, an entrained bed reactor, a transport bed reactor, or amechanical conveying system. The preferred metal for the redox operationinclude but are not limited to Co, Fe, Cu, Ni, Mn, and W. The supportmaterial and the metal are selected such that the metal oxide compositeis not very catalytically active for tar cracking.

FIG. 8 illustrates another embodiment for the integration of a biomassfast pyrolysis and redox process. In this embodiment, metal oxidecomposite does not directly contact the biomass feed, i.e. heat isindirectly provided to the fast pyrolyzer. In this embodiment, the fuelfor the reducer is again the byproducts and char from fast pyrolysis ofbiomass. The reducer reduces the hot metal oxide from the oxidizer:MeO+unwanted fuel from pyrolyzer and refining block→MeO_(y)+CO₂This step is often endothermic, the hot MeO_(y) exiting the reducer at atemperature ranging between 400-750° C.

The reduced MeO_(y) then enters the oxidizer which is preferably anentrained bed, transport bed, or a fast fluidized bed reactor. Theoxidizer is designed similar to a shell and tube heat exchanger withmetal oxide composite and air flowing in the shell side. Air oxidizesMeO_(y) back to MeO_(x).MeO_(y)+O₂→MeO_(x)Significant heat is generated in this step. Meanwhile, high temperatureexhaust air is also generated. The reducer can be either a moving bed ora fluidized bed.

The N₂ rich exhaust air, with a small amount of residual oxygen, can bedirectly used for biomass feeding and conveying in the fast pyrolyzer toprovide the heat. In certain embodiments, an additional combustion stepwith excess amounts of byproduct fuel from the fast pyrolysis stage canbe used to remove the residual oxygen prior to using the hightemperature N₂ rich gas for biomass feeding and conveying.

Pulverized biomass is introduced into the pyrolyzer which is installedinside the oxidizer. The pulverized biomass, carried by the hightemperature gas, is injected in a tangential direction into thepyrolyzer and is conveyed upwards by the high temperature gas in aswirling manner. The centrifugal force causes the biomass to be close tothe pyrolyzer/oxidizer wall through which heat can be transferred to thebiomass for pyrolysis. The pyrolyzer is a fast fluidized bed, entrainedbed, or a dilute transport bed.

Alternatively, the reducer can be integrated with the pyrolyzer toprovide the heat to the pyrolyzer from its outer wall. In both cases,the pyrolyzer is operated at between 300-650° C., the reducer isoperated at between 400-1300° C., and the oxidizer is operated atbetween 450-1350° C.

The performance of the reducer in the redox based reforming/gasificationsub-system is important to the success of the integrated embodiments asshown in FIG. 1, FIG. 2, FIG. 3, FIG. 4, FIG. 6, FIG. 7, and FIG. 8. Inaddition to the designs disclosed in Fan, PCT Application No. WO2007082089; and Fan, PCT Application No. WO 2010037011. improvementshave made in the reducer design for conversion of solid fuels.

FIG. 9 illustrates an improved design of the reducer. In this design,metal oxide composite particles, which are large (0.5-10 mm) and moredense (>1.5 g/mL), are fed from the top of the reducer. The pulverizedbiomass or coal or other solid fuels, which are small (<0.5 mm) and lessdense (<1.5 g/mL) are fed to the bottom section of the reducer. Thepulverized coal or biomass is entrained by the conveying gas and flowsupwards between the gaps of the composite particles while beingconverted. The composite particles move downwards and are reduced beforeexiting the reducer.

The invention claimed is:
 1. A method for producing a liquid fuel from asolid carbonaceous fuel comprising: directly liquefying a first portionof the solid carbonaceous fuel with a hydrogen rich gas stream to formthe liquid fuel; and indirectly gasifying a second portion of the solidcarbonaceous fuel to form separate streams of hydrogen and CO₂ richgases by: reducing metal oxide containing particles in a first reactionzone with the second portion of the solid carbonaceous fuel therebyforming the CO₂ rich gases and reduced metal oxide containing particles;directly sending a first portion of the reduced metal oxide containingparticles from the first reaction zone to a second reaction zone, and asecond portion of the reduced metal oxide containing particles from thefirst reaction zone to a third reaction zone; oxidizing the firstportion of the reduced metal oxide containing particles in the secondreaction zone with steam thereby generating the hydrogen rich gases andat least partially oxidized metal oxide containing particles; sendingthe at least partially oxidized metal oxide containing particles fromthe second reaction zone to the third reaction zone; oxidizing the atleast partially oxidized metal oxide containing particles from thesecond reaction zone and the second portion of the reduced metal oxidecontaining particles from the first reaction zone in the third reactionzone with an oxygen containing gas, thereby generating oxidized metaloxide containing particles; and returning the oxidized metal oxidecontaining particles from the third reaction zone to the first reactionzone; and reacting hydrogen from the hydrogen rich gases with carbondioxide from the CO₂ rich gases in a CO₂ hydrogenation reaction to formsynthetic liquid fuel.
 2. The method of claim 1, where the metal oxidecontaining particle is an iron oxide containing particle.
 3. The methodof claim 1, where the metal oxide containing particles in the firstreaction zone form a packed moving bed.
 4. The method of claim 1, wherethe first reaction zone is operated at a temperature of greater than orequal to 400° C. and less than or equal to 1200° C. and at a pressure ofgreater than or equal to 1.01×10⁵ Pa and less than or equal to 8.10×10⁶Pa.
 5. The method of claim 1, where at least a portion of the carbondioxide from the CO₂ rich gases is sequestered.
 6. The method of claim1, where at least a portion of the steam is generated using heat fromthe CO₂ hydrogenation reaction.